Apparatus for steam-methane reforming

ABSTRACT

Apparatuses for use in plants for processing methane, the apparatuses comprising a plurality of reaction modules each including a plurality of Fischer-Tropsch reactors operable to convert a gaseous mixture including carbon monoxide and hydrogen to a liquid hydrocarbon. Each module may be disconnected and taken away for servicing while allowing the plant to continue to operate. In some of the apparatuses, each Fischer-Tropsch reactor comprises a plurality of metal sheets arranged as a stack to define first and second flow channels for flow of respective fluids, the channels being arranged alternately to ensure good thermal contact between the fluids in the channels.

CROSS-REFERENCE TO RELATED APPLICATIONS

This is a continuation of prior pending U.S. application Ser. No.13/235,784, filed on Sep. 19, 2011 and currently allowed, which is acontinuation of prior pending application Ser. No. 12/081,303, filed onApr. 14, 2008 and issued as U.S. Pat. No. 8,021,633, which is acontinuation of application Ser. No. 10/497,785, filed on Jun. 29, 2004and abandoned.

BRIEF DESCRIPTION OF THE FIGURES

FIG. 1 shows a flow diagram of a chemical process of the invention.

FIG. 2 shows a plan view of a reactor suitable for performing a step ofthe process shown in FIG. 1.

DETAILED DESCRIPTION

This invention relates to a chemical process, and to a plant includingcatalytic reactors suitable for use in performing the process.

A process is described in WO 01/51194 (Accentus plc) in which methane isreacted with steam, to generate carbon monoxide and hydrogen in a firstcatalytic reactor; the resulting gas mixture is then used to performFisher-Tropsch synthesis in a second catalytic reactor. The overallresult is to convert methane to hydrocarbons of higher molecular weight,which are usually liquid under ambient conditions. The two stages of theprocess, steam/methane reforming and Fisher-Tropsch synthesis, requiredifferent catalysts, and catalytic reactors are described for eachstage. The catalytic reactors enable heat to be transferred to or fromthe reacting gases, respectively, as the reactions are respectivelyendothermic and exothermic; the heat required for steam/methanereforming is provided by combustion of methane. A potential problem withthis process is that other reactions may occur in the steam/methanereformer reactor, either to generate carbon dioxide, or to generatecoke. It is suggested that the reformer may incorporate aplatinum/rhodium catalyst, the reaction being performed at 800° C. Thesuggested process relies on a steam/methane ratio that is close to 1:1,as the rhodium catalyst is apparently resistant to coking. An improvedway of performing this process has now been found.

According to the present invention there is provided a process forperforming steam/methane reforming to generate carbon monoxide andhydrogen, wherein the gas mixture is caused to flow through a narrowflow channel between metal sheets separating the flow channel from asource of heat, the flow channel containing a fluid-permeable catalyststructure, the residence time in the channel being less than 0.5 second,and both the average temperature along the channel and the exittemperature of the channel being in the range 750° C. to 900° C.,wherein the steam is supplied at least in part by condensing water vaporfrom combustion of a combustible gas, preferably comprising methane.

The present invention also provides a plant for performing steam/methanereforming that is particularly adapted for use on an oil rig, a floatingplatform or a ship. Under such circumstances space is limited, and theweight of the equipment must be minimized.

Preferably the residence time is less than 0.1 s, but preferably atleast 0.02 s. It is presumed that such short reaction times enable theprocess to operate under non-equilibrium conditions, so that only thosereactions that have comparatively rapid kinetics will occur. It is alsopreferable that the ratio of steam to methane should be in the range 1.2to 2.0, more preferably 1.3 to 1.6, more preferably about 1.4 or 1.5.Under these conditions the proportion of methane that undergoes reactioncan exceed 90%. Furthermore the selectivity in formation of carbonmonoxide rather than carbon dioxide can exceed 85% and even 90%.

The catalytic reactor preferably comprises a plurality of metal sheetsarranged to define first and second flow channels, the channels beingarranged alternately to ensure good thermal contact between the fluidsin them. Appropriate catalysts should be provided in each channel,depending on the required reaction. To ensure the required good thermalcontact, both the first and the second flow channels are preferably lessthan 5 mm deep in the direction normal to the sheets. More preferablyboth the first and the second flow channels are less than 3 mm deep.Corrugated or dimpled foils, metal meshes, or corrugated or pleatedmetal felt sheets may be used as the substrate of the catalyst structurewithin the flow channels to enhance heat transfer. Since good heattransfer is needed for achieving high CO selectivity in thesteam/methane reforming, a preferred structure comprises a metal foilwith a thin coating comprising the catalyst material.

As described in WO 01/51194, such a reactor may be used for performingmethane/steam reforming, the alternate channels containing a methane/airmixture so that the exothermic oxidation reaction provides the necessaryheat for the endothermic reforming reaction. For the oxidation reactionseveral different catalysts may be used, for example palladium, platinumor copper on a ceramic support; for example copper or platinum on analumina support stabilized with lanthanum, cerium or barium, orpalladium on zirconia, or more preferably platinum/palladium on gammaalumina with a metal loading of about 10% by weight (relative to thealumina). This catalyst composition is preferably in a coating ofthickness 20 to 200 μm on a surface in the channel, preferably on thefoil. For the reforming reaction also several different catalysts may beused, for example nickel, platinum, palladium, ruthenium or rhodium,which may be used on ceramic coatings; the preferred catalyst for thereforming reaction is platinum with rhodium as a promoter, on alumina orstabilized alumina. Again the catalyst metal is preferably about 10% byweight compared to the alumina, and is provided as a 10-200 μm coating,preferably 10 to 50 μm. The oxidation reaction may be carried out atsubstantially atmospheric pressure, while although the reformingreaction may be carried out at elevated pressure, for example up to 2MPa (20 atmospheres), operation at atmospheric pressure is preferred, orpossibly slightly elevated pressure for example in the range 0 to 200kPa above atmospheric pressure.

It will be appreciated that the materials of which the reactor are madeare subjected to a severely corrosive atmosphere in use, for example thetemperature may be as high as 900° C., although more typically around800° C. or 850° C. The reactor may be made of a metal such as analuminum-bearing ferritic steel, in particular of the type known asFecralloy which is iron with up to 20% chromium, 0.5-12% aluminum, and0.1-3% yttrium. For example it might comprise iron with 15% chromium, 4%aluminum, and 0.3% yttrium. When this metal is heated in air it forms anadherent oxide coating of alumina which protects the alloy againstfurther oxidation; this oxide layer also protects the alloy againstcorrosion under conditions that prevail within for example a methaneoxidation reactor or a steam/methane reforming reactor. Where this metalis used as a catalyst substrate, and is coated with a ceramic layer intowhich a catalyst material is incorporated, the alumina oxide layer onthe metal is believed to bind with the oxide coating, so ensuring thecatalytic material adheres to the metal substrate.

For some purposes the catalyst metal might instead be deposited directlyonto the adherent oxide coating of the metal (without any ceramiclayer).

The gases produced by the steam/methane reforming process describedabove are preferably then subjected to Fischer-Tropsch synthesis. Thismay be performed using a second such reactor, with a different catalyst.Where excess hydrogen remains, after the Fisher-Tropsch synthesis, thishydrogen is preferably separated from the desired products, and fed backto the combustion flow channels of the steam/methane reforming reactor.Combustion of a mixture of methane and hydrogen with air in thesechannels has been found to give more uniform temperature, and alsoenables the combustion reaction to be initiated more readily when thereactor is cold.

The invention will now be further and more particularly described, byway of example only, and with reference to the accompanying drawings inwhich:

FIG. 1 shows a flow diagram of a chemical process of the invention; and

FIG. 2 shows a plan view of a reactor suitable for performing a step ofthe process shown in FIG. 1.

The invention relates to a chemical process for converting methane tolonger chain hydrocarbons. The first stage involves steam/methanereforming, that is to say the reaction:

H₂O+CH₄→CO+3H₂

This reaction is endothermic, and may be catalyzed by a platinum/rhodiumcatalyst in a first gas flow channel. The heat required to cause thisreaction may be provided by combustion of methane, that is to say:

CH₄+2O₂→CO₂+2H₂O

which is an exothermic reaction, and may be catalyzed by aplatinum/palladium catalyst in an adjacent second gas flow channel. Boththese reactions may take place at atmospheric pressure, althoughalternatively the reforming reaction might take place at an elevatedpressure. The heat generated by the combustion reaction would beconducted through the metal sheet separating the adjacent channels.

The gas mixture produced by the steam/methane reforming can then be usedto perform a Fischer-Tropsch synthesis to generate a longer chainhydrocarbon, that is to say:

nCO+2nH₂→(CH₂)_(n)+nH₂O

which is an exothermic reaction, occurring at an elevated temperature,typically between 200 and 350° C., for example 230° C., and an elevatedpressure typically between 2 MPa and 4 MPa, for example 2.5 MPa, in thepresence of a catalyst such as iron, cobalt or fused magnetite, with apromoter such as potassium. The exact nature of the organic compoundsformed by the reaction depends on the temperature, the pressure, and thecatalyst, as well as the ratio of carbon monoxide to hydrogen. Apreferred catalyst is γ-alumina (as a coating) of surface area 140-230m²/g, with about 35% by weight of cobalt with a ruthenium, platinum,gadolinium or rhenium promoter and a basicity promoter such as ThO₂. Theheat given out by this synthesis reaction may be used to provide atleast part of the heat required by the steam/methane reforming reaction,for example a heat transfer fluid may be used to transfer the heat froma reactor in which the Fischer-Tropsch synthesis is occurring, alsoensuring the temperature in the Fischer-Tropsch reactor remains steady,the heat being used to preheat at least one of the gas streams suppliedto the reforming reactor.

These reactions would be particularly advantageous if they could beperformed at sea, for example on a floating production platform or anoil rig, as they would enable stranded gas resources to be exploited.Stranded gas or associated gas reserves represent an untapped source offuel, but cannot readily be exploited because they are often locatedremotely, and the gas flows may not be large enough to justifyconstruction of a pipeline or a plant to produce liquefied natural gas.Currently such gas is usually flared or re-injected. A plant forproducing liquid hydrocarbon by steam/methane reforming followed byFischer-Tropsch synthesis on land is known, but conventional plant forthis purpose is much too large to conveniently be accommodated on thedeck of a floating structure, and indeed some of the processes employedwould be vulnerable to wave-induced motion. However, by performing suchreactions using compact catalytic reactors, for example as described inPCT/GB2002/004144, which are typically one-tenth the volume of aconventional reactor for the same duty, it is feasible to perform thisprocess on a rig or a floating structure.

It will be appreciated from the equations given above that steam must beprovided to perform the steam/methane reforming reaction, and indeed toensure that the catalysts do not become coated with coke it is,necessary to provide more moles of steam than of methane. Ideally allthe water provided to perform steam/methane reforming could be recoveredafter the Fischer-Tropsch synthesis, but in practice additional water isrequired because of inefficiencies in each stage. For operation on a rigor a floating structure at sea, it will be appreciated that such watercould in principle be provided by a distillation plant fed with seawater, but boiling sea water generates salt, and tends to lead tocorrosion problems; it would be preferable if the steam could beprovided by the chemical process plant itself

Referring now to FIG. 1, the plant and overall chemical process is shownas a flow diagram. The feed gas 4 consists primarily of methane, with asmall percentage (say 10%) of ethane and propane. It may also containcompounds of sulphur that would be detrimental to catalysts. It is firstpassed through a fluidic vortex scrubber 5 in which it flows radiallyinward in counter-current to droplets of a de-sulphurisation liquid.This may for example comprise an aqueous solution of a chelated ferricsalt that reacts with sulphurous compounds and is thereby reduced to theferrous form. The liquid is recirculated by a pump 6 through a fluidicvortex scrubber 7 in which it is contacted by air to regenerate theferric salt and to form a precipitate of sulphur, and then through afilter 8 and a pump 9 back to the scrubber 5. The feed gas 4 is hencede-sulphurised. The vortex scrubbers 5 and 7 are not vulnerable towave-induced motion.

Alternatively the sulphur-contaminated natural gas may be reacted withhydrogen, at a temperature of 200-500° C., over a hydro-desulphurisationcatalyst, to convert mercaptans to H₂S. The gas can then be passedthrough a bed of an adsorbent (such as ZnO so the H₂S reacts to give H₂Oand ZnS). Some adsorbents can be regenerated in situ, producing SO₂.

The de-sulphurised feed gas 10 is passed through a heat exchanger 11 soit is at about 400° C. and is then supplied via a fluidic vortex mixer12 to a first catalytic reactor 14; in the mixer 12 the feed gas ismixed with a stream of steam that is also at about 400° C., the streamsentering the mixer 12 through tangential inlets and following a spiralpath to an axial outlet so they become thoroughly mixed. Both streamsmay be at atmospheric pressure, or for example at a pressure of say 100kPa above atmospheric. The flows are preferably such that the steam;methane molar ratio (at the steam/methane reforming stage) is between1.4 and 1.6, preferably 1.5. The first part of the reactor 14 is apre-reformer 15 with a nickel methanation catalyst at 400° C., in whichthe higher alkanes react with the steam to form methane (and carbonmonoxide); extra steam is required to ensure the desired steam/methaneratio is achieved after this pre-forming stage (this pre-reformer 15would not be required if the feed gas 4 contained substantially nohigher alkanes). An alternative catalyst for the pre-reformer isplatinum/rhodium. The second part of the reactor 14 is a reformer 16with a platinum/rhodium catalyst, in which the methane and steam reactto form carbon monoxide and hydrogen. This reaction may be performed ataround 850° C., as described below.

The heat for the endothermic reactions may be provided by combustion ofmethane over a palladium or platinum catalyst within adjacent gas flowchannels 17 of the reactor 14. The catalyst may incorporate a metalhexaaluminate (such as magnesium hexaaluminate) or more preferablyγ-alumina, with 5-20% (say 10%) by weight palladium/platinum catalyst.The methane/oxygen mixture may be supplied to the channels 17 in stagesalong their length, to ensure combustion occurs throughout their length.The exhaust gases from the combustion channels 17 are passed through asea water-cooled heat exchanger 19 to cause at least part of the watervapor to condense, the remaining gases being released to the atmosphereas exhaust while the liquid water is fed through the duct 24 (seebelow).

The hot mixture of carbon monoxide and hydrogen emerging from thereformer 16 is then quenched by passing through a heat exchanger 18 toprovide the hot steam supplied to the vortex mixer 12, and then throughthe heat exchanger 11 in which it loses heat to the feed gas. Themixture is then further cooled to about 100° C. by passing through aheat exchanger 20 cooled by water. Any water vapor that condenses isseparated from the gas stream into duct 25. The gases are thencompressed through a compressor 22 to a pressure in the range 1.0 MPa to2.5 MPa (10 to 25 atm.).

The stream of high pressure carbon monoxide and hydrogen is thensupplied to a catalytic reactor 26 in which the gases react, undergoingFischer-Tropsch synthesis to form a paraffin or similar compound. Thisreaction is exothermic, preferably taking place at about 230° C., andthe heat generated may be used to preheat the steam supplied to the heatexchanger 18, using a heat exchange fluid such as helium circulatedbetween heat exchange channels in the reactor 26 and a steam generator28. During this synthesis the volume of the gases decreases. Theresulting gases are then passed into a condenser 30 in which theyexchange heat with water initially at 25° C. The higher alkanes (say C5and above) condense as a liquid, as does the water, this mixture ofliquids being passed to a gravity separator 31; the separated higheralkanes can then be removed as the desired product.

The water from the separator 31 is returned via the heat exchangers 28and 18 to the mixer 12. The water from the ducts 24 and 25 is alsocombined with this water stream. The water in the separator 31 may alsocontain alcohols (which may be formed in the Fischer-Tropsch reactor26), so the water may first be steam-stripped to remove such solubleorganic compounds before it is returned to the mixer 12. If water thatcontains alcohols is returned to the mixer 12, the alcohols will bereformed to produce CO, CO_(2 and H) ₂.

Any lower alkanes or methane, and remaining hydrogen, pass through thecondenser 30 and are supplied to a refrigerated condenser 32 in whichthey are cooled to about 5° C. The gases that remain, consistingprimarily of hydrogen with carbon dioxide, methane and ethane, arepassed through a pressure-reducing turbine 33 and fed via a duct 34 intoa storage vessel 35, and hence through a valve 36 into the combustionchannel of the first catalytic reactor 14. The condensed vapors,consisting primarily of propane, butane and water, are passed to agravity separator 37, from which the water is combined with the recycledwater from the separator 31, while the alkanes are recycled to the feedline 10 so as to be fed into the pre-reformer 15. As indicated by thebroken line, electricity generated by the turbine 33 may be used to helpdrive the compressor 22.

When used in this fashion the overall result of the processes is thatmethane is converted to higher molecular weight hydrocarbons which aretypically liquids at ambient temperatures. The processes may be used atan oil or gas well to convert methane gas into a liquid hydrocarbonwhich is easier to transport.

From the steam/methane reforming reaction given above one would expectthat the appropriate ratio between steam and methane would be 1 to 1.However, at that ratio there is a significant risk of coking, and a riskthat a significant proportion of the methane will not undergo thereaction. Increasing the proportion of steam increases the proportion ofmethane that reacts, and decreases the risk of coking, although if theproportion of steam is too high then there is an increased likelihood ofcarbon dioxide formation. It has been found that operating with asteam/methane ratio of between 1.3 and 1.6, preferably 1.4 or 1.5,combined with short residence times that are preferably no more than 100ms, gives both high selectivity for carbon monoxide formation and also avery high proportion of methane undergoing reaction. The flow ratesthrough the reformer 16 are preferably such that the residence time isin the range 20 to 100 ms, more preferably about 50 ms. The averagetemperature along each channel in the reformer 16 is above 750° C.,preferably between 800° C. and 900° C.

Such a short residence time enables the reactor 16 to operate under whatappears to be a non-equilibrium condition. The competing reactionbetween carbon monoxide and steam to form the unwanted products carbondioxide and hydrogen has slower kinetics than the steam/methanereforming reaction to form carbon monoxide and hydrogen; and in thereforming reaction the reverse process has slower kinetics than theforward reaction. The short residence time allows insufficient time forthe slower reactions to reach equilibrium. Under these circumstances theproportion of methane undergoing reaction may exceed 90%, and theselectivity for carbon monoxide can exceed 90%.

Experimental measurements have been made, passing a preheatedsteam/methane mixture (ratio 1.5) through a multichannel reactor similarto that described below with reference to FIG. 2, with a residence timeof 50 ms. The temperature was measured near the inlet and near the exitfrom a channel, and at other intermediate positions to enable a meanvalue to be calculated. Results have been obtained as in the Table:

Inlet Temp Exit Temp Mean Temp CO Selectivity CH4 Conversion (° C.) (°C.) (° C.) (%) (%) 726 853 810 92.5 93.8 731 866 815 93.0 94.2

Selectivity for CO production is enhanced by operating with an exittemperature above 800° C., more preferably above 850° C. The performanceof the reactor can also be improved by a pre-treatment, heating thereactor to about 850° C. in the presence of hydrogen, as this improvessubsequent catalyst activity.

As indicated above, the ideal hydrogen to carbon monoxide stoichiometricratio to feed to the Fischer-Tropsch synthesis reactor would be about 2moles hydrogen to 1 mole carbon monoxide. This ratio cannot readily beobtained by steam/methane reforming: as discussed above, at asteam/methane ratio of 1.0 the resulting gas mixture has a hydrogen tocarbon monoxide ratio 3 to 1, and at the elevated steam/methane ratiosthat must be adopted to avoid coking the hydrogen to carbon monoxideratio is above 3, and may be as high as 4 to 1. Consequently, after theFischer-Tropsch synthesis reaction has occurred there will be an excessof hydrogen that remains Feeding this gas into the combustion channel ofthe reactor 14 has been found to give a more uniform temperaturedistribution, and also enables the combustion reaction to be initiatedmore readily when the reactor is cold (as catalytic combustion can thenoccur at a temperature as low as 15 or 20° C.). The overall thermalefficiency of the process is improved, the amount of methane feddirectly to the combustion channels is decreased, and the emission ofcarbon dioxide to the environment is also reduced.

Referring now to FIG. 2, a reactor 40 (suitable for example forsteam/methane reforming as reactor 14) comprises a stack of Fecralloysteel plates 41, each plate being generally rectangular, 450 mm long and150 mm wide and 3 mm thick, these dimensions being given only by way ofexample. On the upper surface of each such plate 41 are rectangulargrooves 42 of depth 2 mm separated by lands 43 (eight such grooves beingshown), but there are three different arrangements of the grooves 42. Inthe plate 41 shown in the drawing the grooves 42 extend diagonally at anangle of 45° to the longitudinal axis of the plate 41, from top left tobottom right as shown. In a second type of plate 41 the grooves 42 a (asindicated by broken lines) follow a mirror image pattern, extendingdiagonally at 45° from bottom left to top right as shown. In a thirdtype of plate 41 the grooves 42 b (as indicated by chain dotted lines)extend parallel to the longitudinal axis.

The plates 41 are assembled in a stack, with each of the third type ofplate 41 (with the longitudinal grooves 42 b) being between a plate withdiagonal grooves 42 and a plate with mirror image diagonal grooves 42 a,and after assembling many plates 41 the stack is completed with a blankrectangular plate. The plates 41 are compressed together and diffusionbonded, so they are sealed to each other. Corrugated Fecralloy alloyfoils 44 (only one is shown) 50 μm thick coated with a ceramic coatingof thickness 15 μm containing a catalyst material, of appropriate shapesand with corrugations 2 mm high, can be slid into each such groove 42,42 a and 42 b. The corrugations extend parallel to the flow direction ineach case.

Header chambers 46 are welded to the stack along each side, each header46 defining three compartments by virtue of two fins 47 that are alsowelded to the stack. The fins 47 are one third of the way along thelength of the stack from each end, and coincide with a land 43 (or aportion of the plates with no groove) in each plate 41 with diagonalgrooves 42 or 42 a. Gas flow headers 48 in the form of rectangular capsare then welded onto the stack at each end, communicating with thelongitudinal grooves 41 b. In a modification (not shown), in place ofeach three-compartment header 46 there might instead be three adjacentheader chambers, each being a rectangular cap like the headers 48.

In use of the reactor 40 for steam/methane reforming, a steam/methanemixture is supplied to the header 48 at one end (the right hand end asshown), and the resulting mixture of hydrogen and carbon monoxideemerges through the header 48 at the other end. Methane/air mixture issupplied to the compartments of both headers 46 at the other end (theleft hand end as shown), and so exhaust gas from the combustion processemerges through the compartments of both headers 46 at the right handend as shown. The flow path for the mixture supplied to the top-leftheader compartment (as shown) is through the diagonal grooves 42 intothe bottom-middle header compartment, and then to flow through thediagonal grooves 42 a in other plates in the stack into the top-rightheader compartment. Hence the gas flows are at least partiallycounter-current, so that the hottest region in the combustion channels,which is near the inlet to those channels, is closest to the outlet forthe steam/methane reforming reaction.

The headers 46 and 48 each comprise a simple rectangular cap sealedaround its periphery to the outside of the stack so as to cover part ofone face of the stack. They may be welded onto the outside of the stack.Alternatively, if neither of the gas flows are at elevated pressures, itmay be adequate to clamp the header chambers 46, 48 onto the outside ofthe stack. In either case it will be appreciated that after a period ofuse, if the catalyst in either or both of the channels has become spent,then the headers 46 and 48 may be removed or cut off and thecorresponding catalyst-carrying foils 44 removed and replaced. Theheaders 46, 48 can then be re-attached.

It will be understood that the type and thickness of ceramic on thecorrugated foils 44 in the flow channels may be different in successiveplates 41 in the stack, and that the catalyst materials may differ also.For example the ceramic might comprise alumina in one of the gas flowchannels, and zirconia in the other gas flow channels. The reactor 40formed from the plates 41 might also be suitable for performingFischer-Tropsch synthesis. Because the plates 41 forming the stack arebonded together the gas flow channels are gas tight (apart fromcommunication with headers 46 or 48), and the dimensions of the plates41 and grooves 42 are such that pressures in the alternate gas flowchannels may be considerably different. Furthermore the pitch or patternof the corrugated foils 44 may vary along a reactor channel 42 to adjustcatalytic activity, and hence provide for control over the temperaturesor reaction rates at different points in the reactor 40. The corrugatedfoils 44 may also be shaped, for example with perforations, to promotemixing of the fluid within the channels 42. Furthermore parts of thefoils 44 may be devoid of catalyst.

It will be appreciated that the plates forming the stack may be of adifferent size (typically in the range 400-1200 mm long, 150-600 mmwide), and that the diagonal grooves 42 and 42 a may have a differentorientation (typically between 30° and 90°), for example the platesmight be 800 mm by 400 mm, and the grooves be at about 56° to thelongitudinal axis (if there are three header compartments along eachside) or at about 63° (if there are four header compartments). In everycase the headers ensure the fluid in those sets of channels follows aserpentine path along the reactor.

In a modification to the reactor 40, the foils 44 are again of Fecralloymaterial, but the catalyst material is deposited directly onto the oxidelayer of the Fecralloy.

Particularly where the reactor 40 is to be used for Fischer-Tropschsynthesis, the gas flow channels 42 for that reaction may decrease inwidth, and possibly also depth, along their length, so as to vary thefluid flow conditions, and the heat or mass transfer coefficients.During the synthesis reaction the gas volume decreases, and byappropriate tapering of the channels 42 the gas velocity may bemaintained as the reaction proceeds. Furthermore the pitch or pattern ofthe corrugated foils 44 may vary along a reactor channel 42 to adjustcatalytic activity, and hence provide for control over the temperaturesor reaction rates at different points in the reactor 40.

When a reactor such as the reactor 40 is used for reactions betweengases that generate gaseous products then the orientation of thechannels is not of concern. However if a product may be a liquid, it maybe preferable to arrange the reactor 40 so that the flow paths for thisreaction slope downwardly, to ensure that any liquid that is formed willdrain out of the channels 42. With the gas flowing along thecorrugations in the foils 44, any liquid tends to be entrained, sominimizing liquid build-up on the surface of the catalyst.

It will be appreciated that the although the heat for the steam/methanereforming reaction may be provided by catalytic combustion in adjacentchannels (as described above), as an alternative the combustion may takeplace in an external burner (such as a laminar flow burner), the veryhot exhaust gases at about 900 or 1000° C. being passed through thesecond gas flow channels of the reactor 14 of FIG. 1 in counter-currentto the methane flow; this can enable the reacting gases in the reformer16 to reach a final temperature of as much as 900° C. In this case it isnot essential to provide the foils in the combustion channels withceramic coating or catalyst, but such foils would nevertheless enhanceheat transfer between the second gas flow channel carrying the hotexhaust gas and the reactants in the pre-reformer and reformer channels,by transferring heat to the separating plates 41. In the combustionchannels of the catalytic reactor 14, if catalytic combustion is used togenerate the heat (as indicated), the combustion catalyst may itself becoated with a thin porous inert ceramic layer, so as to restrict thecontact of the gas mixture with the catalyst and so restrict thereaction rate particularly at the start of the channel.

Particularly where hydrogen is unavailable, it may be desirable toprovide electrical heating by passing an electric current directlythrough the plates forming the reactor. This may be used initially toraise the temperature for example of the reforming reactor 14 to say400° C. before supplying gases, to ensure catalytic combustion occurs.Such electrical heating may also be used during operation to adjust thereactor temperature. Electrical heating may also be used in the vicinityof the outlet from the reactor 14 to ensure that a temperature of say900° C. is reached by the gases undergoing the reforming reaction.

As mentioned above the reactor may differ in size or shape from thatshown in FIG. 2. A single such plate might instead for example be 1.0 mby 0.5 m. The stack forming a reactor might be 0.8 m thick. Several suchreactors may be combined into a reaction module, for example ten suchreactors might form a module provided with pipes interconnected so thegas flows are in parallel through all the reactors in the module. Such amodule may be small enough to be transported in an ISO structural frame,and yet have sufficient capacity to produce synthesis gas equivalent to1000 barrels per day of synthetic oil.

A practical plant may need to include several such modules, all beingoperated with gas flows in parallel, although it may not be necessary toreplicate the other features (e.g. heat exchangers 18, 11, 20 etc. andseparators 31 and 35) of the plant. Thus there might be say six or tenmodules made up of the reactors 14, and the same number of modules ofFischer-Tropsch reactors 26. This has the benefit that if the catalystsin one module need to be replaced, that module may be disconnected andtaken away for servicing, while allowing the plant to continue tooperate at only slightly reduced capacity.

1. An apparatus for use in a plant for processing methane, the apparatuscomprising a plurality of reaction modules each including a plurality ofFischer-Tropsch reactors operable to convert a gaseous mixture includingcarbon monoxide and hydrogen to a liquid hydrocarbon, wherein eachmodule may be disconnected and taken away for servicing, whilst allowingthe plant to continue to operate.
 2. An apparatus according to claim 1wherein each Fischer-Tropsch reactor comprises a plurality of metalsheets arranged as a stack to define first and second flow channels forflow of respective fluids, the channels being arranged alternately toensure good thermal contact between the fluids in the channels.
 3. Anapparatus according to claim 2 wherein both the first flow channels andthe second flow channels are less than 5 mm deep in a direction normalto the center plane of the sheets
 4. The apparatus according to claim 1,wherein at least one of the plurality of Fisher-Tropsch reactorsincludes a removable catalyst.
 5. The apparatus according to claim 1,wherein at least one of the plurality of Fisher-Tropsch reactorsincludes a catalyst selected from the group comprising iron, cobalt andfused magnetite.
 6. The apparatus according to claim 5, wherein thecatalyst is supported on γ-alumina.
 7. The apparatus according to claim6, wherein the γ-alumina has a surface area of 140-230 m2/g.
 8. Theapparatus according to claim 1, wherein at least one of the plurality ofFisher-Tropsch reactors includes a catalyst comprising about 35% byweight of cobalt.
 9. The apparatus according to claim 5, wherein thecatalyst is provided with a promoter selected from the group comprisingruthenium, platinum, gadolinium and rhenium.
 10. The apparatus accordingto claim 1, wherein the reaction module is small enough to betransported in an ISO structural frame.
 11. The apparatus according toclaim 1, wherein the plurality of Fisher-Tropsch reactors are arrangedsuch that the gaseous mixture may flow in parallel through the pluralityof Fisher-Tropsch reactors.